Method for the dehydrogenation of hydrocarbons

ABSTRACT

In a process for the heterogeneously catalyzed dehydrogenation in one or more reaction zones of one or more dehydrogenatable C 2 -C 30 -hydrocarbons in a reaction gas mixture comprising them, with at least part of the heat of dehydrogenation required being generated directly in the reaction gas mixture in at least one reaction zone by combustion of hydrogen, the hydrocarbon or hydrocarbons and/or carbon in the presence of an oxygen-containing gas, the reaction gas mixture comprising the dehydrogenatable hydrocarbon or hydrocarbons is brought into contact with a Lewis-acid dehydrogenation catalyst which has essentially no Bronsted acidity.

[0001] The present invention relates to a process for theheterogeneously catalyzed dehydrogenation of dehydrogenatableC₂-C₃₀-hydrocarbons.

[0002] Dehydrogenated hydrocarbons are required in large quantities asstarting materials for numerous industrial processes. For example,dehydrogenated hydrocarbons are used in the production of detergents,antiknock gasoline and pharmaceutical products. Likewise, numerousplastics are produced by polymerization of olefins.

[0003] For example, acrylonitrile, acrylic acid or C₄ oxo alcohols areprepared from propylene. Propylene is at present produced predominantlyby steam cracking or catalytic cracking of suitable hydrocarbons orhydrocarbon mixtures such as naphtha.

[0004] Propylene can also be prepared by heterogeneously catalyzeddehydrogenation of propane.

[0005] To obtain acceptable conversions in heterogeneously catalyzeddehydrogenations even on a single pass through the reactor, relativelyhigh reaction temperatures generally have to be employed. Typicalreaction temperatures for gas-phase dehydrogenations are from 300 to700° C. In general, one molecule of hydrogen is produced per molecule ofhydrocarbon.

[0006] The dehydrogenation of hydrocarbons proceeds endothermically. Theheat of dehydrogenation necessary for obtaining a desired conversion hasto be introduced into the reaction gas either beforehand and/or duringthe course of the catalytic dehydrogenation. In most knowndehydrogenation processes, the heat of dehydrogenation is generatedoutside the reactor and is introduced into the reaction gas from theoutside. However, this requires complicated reactor and process conceptsand leads, particularly at high conversions, to steep temperaturegradients in the reactor, which is accompanied by the risk of increasedbyproduct formation. Thus, for example, a plurality of adiabaticcatalyst beds can be arranged in annular gap reactors connected inseries. The reaction gas mixture is superheated by means of heatexchangers on its way from one catalyst bed to the next catalyst bed andcools down again during passage through the subsequent reactor. Toachieve high conversions using such a reactor concept, it is necessaryeither to increase the number of reactors connected in series or toincrease the reactor inlet temperature of the gas mixture. The resultingsuperheating inevitably leads to increased byproduct formation due tocracking reactions. Another known method is to arrange the catalyst bedin a tube reactor and to generate the heat of dehydrogenation by burningcombustible gases outside the tube reactor and introduce it into theinterior of the reactor via the tube wall. In these reactors, highconversions lead to steep temperature gradients between the wall and theinterior of the reation tube.

[0007] An alternative is to generate the heat of dehydrogenationdirectly in the reaction gas mixture of the dehydrogenation by oxidationof hydrogen formed in the dehydrogenation or additionally fed in or ofhydrocarbons present in the reaction gas mixture by means of oxygen. Forthis purpose, an oxygen-containing gas and possibly hydrogen is/areadded to the reaction gas mixture either upstream of the first catalystbed or upstream of subsequent catalyst beds. The heat of reactionliberated in the oxidation also prevents high temperature gradients inthe reactor at high conversions. At the same time, a very simple processconcept is realized by omission of the indirect heating of the reactor.

[0008] U.S. Pat. No. 4,788,371 describes a process for steamdehydrogenation of dehydrogenatable hydrocarbons in the gas phasecombined with oxidative reheating of the intermediates, with the samecatalyst being used for the selective oxidation of hydrogen and thesteam dehydrogenation. Here, hydrogen can be introduced as co-feed. Thecatalyst used comprises a noble metal of group VIII, an alkali metal anda further metal selected from the group consisting of B, Ga, In, Ge, Snand Pb on an inorganic oxide support such as aluminum oxide. The processcan be carried out in one or more stages in a fixed or moving bed.

[0009] WO 94/29021 describes a catalyst which comprises a supportconsisting essentially of a mixed oxide of magnesium and aluminumMg(Al)O and also a noble metal of group VIII, preferably platinum, ametal of group IVA, preferably tin, and possibly an alkali metal,preferably cesium. The catalyst is used in the dehydrogenation ofhydrocarbons, which can be carried out in the presence of oxygen.

[0010] U.S. Pat. No. 5,733,518 describes a process for the selectiveoxidation of hydrogen by oxygen in the presence of hydrocarbons such asn-butane over a catalyst comprising a phosphate of germanium, tin, lead,arsenic, antimony or bismuth, preferably tin. The combustion of thehydrogen generates, in at least one reaction zone, the heat of reactionnecessary for the endothermic dehydrogenation.

[0011] EP-A 0 838 534 describes a catalyst for the steam-freehydrogenation of alkanes, in particular isobutane, in the presence ofoxygen. The catalyst used comprises a platinum group metal applied to asupport comprising tin oxide/zirconium oxide and having a tin content ofat least 10%. The oxygen content of the feed stream for thedehydrogenation is calculated so that the quantity of heat generated bythe combustion reaction of hydrogen and oxygen is equal to the quantityof heat required for the dehydrogenation.

[0012] WO 96/33151 describes a process for the dehydrogenation of aC₂-C₅-alkane in the absence of oxygen over a dehydrogenation catalystcomprising Cr, Mo, Ga, Zn or a group VIII metal with simultaneousoxidation of the resulting hydrogen over a reducible metal oxide, e.g.an oxide of Bi, In, Sb, Zn, T1, Pb or Te. The dehydrogenation has to beinterrupted at regular intervals in order to reoxidize the reduced oxideby means of an oxygen source. U.S. Pat. No. 5,430,209 describes acorresponding process in which the dehydrogenation step and theoxidation step proceed sequentially and the associated catalysts areseparated physically from one another. Catalsyts used for the selectiveoxidation of hydrogen are oxides of Bi, Sb and Te and also their mixedoxides.

[0013] Finally, WO 96/33150 describes a process in which a C₂-C₅-alkaneis dehydrogenated over a dehydrogenation catalyst in a first stage, theoutput gas from the dehydrogenation stage is mixed with oxygen and, in asecond stage, passed over an oxidation catalyst, preferably Bi₂O₃, so asto selectively oxidize the hydrogen formed to water, and, in a thirdstage, the output gas from the second stage is again passed over adehydrogenation catalyst.

[0014] The catalyst system used has to meet demanding requirements inrespect of achievable alkane conversion, selectivity to formation ofalkenes, mechanical stability, thermal stability, carbonizationbehavior, deactivation behavior, regenerability, stability in thepresence of oxygen and insensitivity to catalyst poisons such as CO,sulfur- and chlorine-containing compounds, alkynes, etc., and economics.

[0015] The catalysts of the prior art do not meet these requirements,particularly in respect of the achievable conversions and selectivities,operating lives and regenerability, to a satisfactory extent.

[0016] It is an object of the present invention to provide a process forthe dehydrogenation of hydrocarbons which ensures high conversions,space-time yields and selectivities.

[0017] We have found that this object is achieved by a process for theheterogeneously catalyzed dehydrogenation in one or more reaction zonesof one or more dehydrogenatable C₂-C₃₀-hydrocarbons in a reaction gasmixture comprising them, with at least part of the heat ofdehydrogenation required being generated directly in the reaction gasmixture in at least one reaction zone by combustion of hydrogen, thehydrocarbon or hydrocarbons and/or carbon in the presence of anoxygen-containing gas, wherein the reaction gas mixture comprising thedehydrogenatable hydrocarbon or hydrocarbons is brought into contactwith a Lewis-acid dehydrogenation catalyst which has essentially noBronsted acidity.

[0018] The dehydrogenation catalyst used according to the presentinvention has essentially no Brönsted acidity, but a high Lewis acidity.The determination of the Lewis and Brönsted acidities of thedehydrogenation catalysts is carried out by adsorption of pyridine asbasic probe molecule on the activated catalyst with subsequentquantitative FT-IR-spectrometric determination of the Brönsted- andLewis-specific adsorbates. This method makes use of the fact that theadsorbed probe molecules give different IR spectra depending on whetherthey are bound to a Bronsted center or a Lewis center. At the Bronstedcenter, proton transfer takes place to form a local ion pair with thepyridinium ion as cation. The adsorbed pyridinium ion displays aBrönsted-specific absorption band at 1545 cm⁻¹ in the IR spectrum. Atthe Lewis center, on the other hand, the probe molecule pyridine iscoordinated via its free electron pair on the ring nitrogen to theelectron-deficient center. This results in an IR spectrum different tothat of the Brönsted adsorbate. The Lewis band is found at 1440 cm⁻¹.Quantitative evaluation of the Brönsted and Lewis bands enables theBrönsted and Lewis centers to be determined separately. The bandassignment is based on the work of Turkevich (C. H. Kline, J. Turkevich:J. Chem. Phys. 12, 300 (1994)).

[0019] The assignment of the resulting pyridine bands in the FT-IRspectrum is as follows:

[0020] Lewis (L): 1440 cm⁻¹

[0021] Brönsted (B): 1545 cm⁻¹

[0022] Control band B+L: 1490 cm⁻¹

[0023] Physisorbed pyridine: 1590 cm⁻¹ (additionally 1440 cm⁻¹)

[0024] The transmission cell used for the measurements is areconstruction of the prototype of Gallei and Schadow (E. Gallei et al.:Rev. Sci. Instrur. 45 (12), 1504 (1976)). The cell comprises a stainlesssteel body with parallel IR-transparent windows made of CaF₂. Theintrinsic absorption of the windows makes it possible to measure only ina spectral range of about 1200-4000 cm⁻¹. A circuit for cooling orheating fluid is provided in the cell body. In the lid of the cell thereis a solid, plate-shaped sample holder with built-in cartridge heating(400° C.). The self-supporting sample compound is laid in an annulardouble template and screwed into the heating plate, and the cell lid isscrewed onto the cell body. The measurement cell can be evacuated to10⁻⁵-10⁻⁶ mbar.

[0025] For sample preparation, the catalyst material is ground finely ina mortar and pressed between two stainless steel plates with micaunderlays in a film press at a pressing pressure of 50 kN to give aself-supporting wafer. The thickness depends on the intrinsic IRabsorption of the material and is typically in the range from 30 to 100μm. Pellets having a diameter of about 5 mm are cut from the wafer.

[0026] The activation of the sample in the measurement cell is carriedout in air at 390° C. After heating, the cell is evacuated to 10⁻⁵-10⁻⁶mbar. It is then cooled to the gas treatment temperature of 80° C. undera high vacuum.

[0027] The sample is subsequently treated with gaseous pyridine at apressure which can be from 10⁻² to 3 mbar. Control spectra of the samplewith adsorbate are recorded until a steady adsorption state has beenestablished at the gas treatment pressure concerned. The cell issubsequently evacuated to a high vacuum (10⁻⁵ mbar). This removesphysisorbates. After evacuation is complete, adsorbate spectra arerecorded.

[0028] To determine the Lewis and Brönsted acidities, the intensities ofthe bands at 1440 cm⁻¹ and 1545 cm⁻¹ obtained at a particular thicknessof the sample and a set equilibrium pressure of pyridine are evaluatedin comparison with one another. If no band at 1545 cm⁻¹ is discernible(no Brönsted acidity), the band at 1490 cm⁻¹ can also be employed fordetermining the Lewis acidity.

[0029] The measured absorbances are expressed as a ratio to thethickness of the sample (in integrated extinction units (IEE) per μm ofthickness). The single-beam spectrum of the sample which has not beentreated with pyridine gas (cooled to 80° C.) under high vacuum serves asbackground of the adsorbate spectra. This completely balances out matrixbands.

[0030] 1 AU corresponds to one thousand times the measured absorbance(reported in integrated extinction units IEE) divided by the thicknessof the sample (in μm) obtained in the determination of the Lewis andBronsted acidities of the dehydrogenation catalysts using the probe gaspyridine.

[0031] The dehydrogenation catalysts used according to the presentinvention generally have no detectable Bronsted acidity, i.e. theirBronsted acidity is less than 0.1 AU. However, they have a high Lewisacidity. The Lewis acidity of the dehydrogenation catalysts is generallygreater than 1 AU, preferably greater than 3 AU, particularly preferablygreater than 6 AU.

[0032] The dehydrogenation catalysts used according to the presentinvention generally comprise a support and an active composition. Thesupport comprises a heat-resistant oxide or mixed oxide. Thedehydrogenation catalyst preferably comprises a metal oxide selectedfrom the group consisting of zirconium dioxide, zinc oxide, aluminumoxide, silicon dioxide, titanium dioxide, magnesium oxide, lanthanumoxide, cerium oxide and mixtures thereof as support. Preferred supportsare zirconium dioxide and/or silicon dioxide; particular preference isgiven to mixtures of zirconium dioxide and silicon dioxide.

[0033] The active composition of the dehydrogenation catalyst usedaccording to the present invention generally comprises one or moreelements of transition group VIII, preferably platinum and/or palladium,particularly preferably platinum. In addition, the dehydrogenationcatalyst can further comprise one or more elements of main groups Iand/or II, preferably potassium and/or cesium. The dehydrogenationcatalyst may also further comprise one or more elements of transitiongroup III including the lanthanides and actinides, preferably lanthanumand/or cerium. Finally, the dehydrogenation catalyst can furthercomprise one or more elements of main groups III and/or IV, preferablyone or more elements selected from the group consisting of boron,gallium, silicon, germanium, tin and lead, particularly preferably tin.

[0034] In a preferred embodiment, the dehydrogenation catalyst comprisesat least one element of transition group VIII, at least one element ofmain groups I and/or II, at least one element of main groups III and/orIV and at least one element of transition group III including thelanthanides and actinides.

[0035] To produce the dehydrogenation catalysts used according to thepresent invention, it is possible to use precursors of oxides ofzirconium, silicon, aluminum, titanium, magnesium, lanthanum or ceriumwhich can be converted into the oxides by calcination. These can beproduced by known methods, for example by the sol-gel process,precipitation of salts, dehydration of the corresponding acids, drymixing, slurrying or spray drying. For example, to produce aZrO₂.Al₂O₃.SiO₂ mixed oxide, a water-rich zirconium oxide of the formulaZrO₂.xH₂O can firstly be prepared by precipitation of a suitablezirconium-containing precursor. Suitable precursors of zirconium are,for example, Zr(NO₃)₄, ZrOCl₂ or ZrC₄. The precipitation itself iscarried out by addition of a base such as NaOH, KOH, Na₂CO₃ or NH₃ andis described, for example, in EP-A 0 849 224.

[0036] To produce a ZrO₂.SiO₂ mixed oxide, the zirconium-containingprecursor obtained above can be mixed with a silicon-containingprecursor. Well-suited precursors of SiO₂ are, for example,water-containing sols of SiO₂ such as Ludox™. The two components can bemixed, for example, by simple mechanical mixing or by spray drying in aspray dryer.

[0037] To produce a ZrO₂.SiO₂.A1 ₂O₃ mixed oxide, the SiO₂.ZrO₂ powdermixture obtained as described above can be admixed with analuminum-containing precursor. This can be achieved, for example, bysimple mechanical mixing in a kneader. However, the ZrO₂.SiO₂.Al₂O₃mixed oxide can also be produced in a single step by dry mixing theindividual precursors.

[0038] The supports for the dehydrogenation catalysts used according tothe present invention have, inter alia, the advantage that they caneasily be shaped. For this purpose, the powder mixture obtained isadmixed with a concentrated acid in a kneader and then converted into ashaped body, e.g. by means of a ram extruder or a screw extruder.

[0039] The dehydrogenation catalysts used according to the presentinvention have, in particular embodiments, a defined pore structure.When using mixed oxides, it is possible to influence the pore structurein a targeted manner. The particle sizes of the various precursorsinfluence the pore structure. Thus, for example, macropores can begenerated in the microstructure by use of Al₂O₃ having a low loss onignition and a defined particle size distribution. In this context, theuse of Al₂O₃ having a loss on ignition of about 3% (e.g. Puralox®) hasbeen found to be useful.

[0040] A further possible way of producing supports having specific poreradius distributions for the dehydrogenation catalysts used according tothe present invention is to add various polymers during production ofthe support and subsequently to remove them completely or partly bycalcination so as to form pores in defined pore ranges. Mixing thepolymers and the oxide precursors can be carried out, for example, bysimple mechanical mixing or by spray drying in a spray dryer.

[0041] The use of PVP (polyvinylpyrrolidone) has been found to beparticularly useful for producing supports having a bimodal pore radiusdistribution. If PVP is added to one or more oxide precursors of oxidesof the elements Zr, Ti, Al or Si in a production step, macropores havingsizes in the range from 200 to 5000 nm are formed after calcination. Afurther advantage of the use of PVP is that it makes the support easierto shape. Thus, extrudates having good mechanical properties can readilybe produced from freshly precipitated hydrous ZrO₂.xH₂O which haspreviously been dried at 120° C. when PVP and formic acid are added,even without further oxide precursors.

[0042] The calcination of the supports for the dehydrogenation catalystsused according to the present invention is advantageously carried outafter application of the active components and is carried out at from400 to 1000° C., preferably from 500 to 700° C., particularly preferablyfrom 550 to 650° C. and in particular at from 560 to 620° C. Thecalcination temperature should usually be at least as high as thereaction temperature of the dehydrogenation in which the dehydrogenationcatalysts are used according to the present invention.

[0043] The supports of the dehydrogenation catalysts used according tothe present invention generally have high BET surface areas aftercalcination. The BET surface areas are generally greater than 40 m²/g,preferably greater than 50 m²/g, particularly preferably greater than 70m²/g. The pore volume of the dehydrogenation catalysts used according tothe present invention is usually from 0.2 to 0.6 ml/g, preferably from0.25 to 0.5 ml/g. The mean pore diameter of the dehydrogenationcatalysts used according to the present invention, which can bedetermined by Hg porosimetry, is from 3 to 20 nm, preferably from 4 to15 nm.

[0044] Furthermore, the dehydrogenation catalysts used according to thepresent invention have a bimodal pore radius distribution. The poreshave sizes in the range up to 20 nm and in the range from 40 to 5000 nm.These pores all together make up at least 70% of the total pore volumeof the dehydrogenation catalyst. The proportion of pores smaller than 20nm is generally in the range from 20 to 60%, while the proportion ofpores in the range from 40 to 5000 nm is generally likewise from 20 to60%.

[0045] The dehydrogenation-active component, which is usually a metal oftransition group VIII, is generally applied by impregnation with asuitable metal salt precursor. Instead of impregnation, thedehydrogenation-active component can also be applied by other methods,for example spraying the metal salt precursor onto the support. Suitablemetal salt precursors are, for example, the nitrates, acetates andchlorides of the corresponding metals; complex anions of the metals usedare also possible. Preference is given to using platinum as H₂PtCl₆ orPt(NO₃)₂. Suitable solvents for the metal salt precursors include bothwater and organic solvents. Particularly useful solvents are water andlower alcohols such as methanol and ethanol.

[0046] When using noble metals as dehydrogenation-active components,suitable precursors also include the corresponding noble metal solswhich can be prepared by one of the known methods, for example byreduction of a metal salt using a reducing agent in the presence of astabilizer such as PVP. The method of preparation is dealt withcomprehensively in the German Patent Application DE 195 00 366.

[0047] The amount of a noble metal present as dehydrogenation-activecomponent in the dehydrogenation catalysts used according to the presentinvention is from 0 to 5% by weight, preferably from 0.05 to 1% byweight, particularly preferably from 0.05 to 0.5% by weight.

[0048] The further components of the active composition can be appliedeither during production of the support, for example by coprecipitation,or subsequently, for example by impregnating the support with suitableprecursor compounds. Precursor compounds used are generally compoundswhich can be converted into the corresponding oxides by calcination.Suitable precursors are, for example, hydroxides, carbonates, oxalates,acetates, chlorides or mixed hydroxycarbonates of the correspondingmetals.

[0049] In advantageous embodiments, the active composition furthercomprises the following additional components:

[0050] at least one element of main group I or II, preferably cesiumand/or potassium, in an amount of from 0 to 20% by weight, preferablyfrom 0.1 to 15% by weight, particularly preferably from 0.2 to 10% byweight;

[0051] at least one element of transition group III including thelanthanides and actinides, preferably lanthanum and/or cerium, in anamount of from 0 to 20% by weight, preferably from 0.1 to 15% by weight,particularly preferably from 0.2 to 10% by weight;

[0052] at least one element of main groups III and IV, preferably tin,in an amount of from 0 to 10% by weight.

[0053] The dehydrogenation catalyst is preferably halogen-free.

[0054] The dehydrogenation catalyst can be used in the form of a fixedbed in the reactor or, for example, in the form of a fluidized bed andcan have an appropriate shape. Suitable shapes are, for example,granules, pellets, monoliths, spheres or extrudates (rods, wagon wheels,stars, rings).

[0055] As dehydrogenatable hydrocarbons, it is possible to useparaffins, alkylaromatics, naphthenes or olefins having from 2 to 30carbon atoms. The process is particularly useful for the dehydrogenationof straight-chain or branched hydrocarbons having a chain length of from2 to 15 carbon atoms, preferably from 2 to 5 carbon atoms. Examples areethane, propane, n-butane, isobutane, n-pentane, isopentane, n-hexane,n-heptane, n-octane, n-nonane, n-decane, n-undecane, ndodecane,n-tridecane, n-tetradecane and n-pentadecane. The particularly preferredhydrocarbon is propane. In the further description of the invention, thediscussion will frequently concern this particularly preferred case ofpropane dehydrogenation, but the corresponding features applyanalogously to other dehydrogenatable hydrocarbons.

[0056] Since the dehydrogenation reaction is accompanied by an increasein volume, the conversion can be increased by lowering the partialpressure of the reactants. This can be achived in a simple manner by,for example, dehydrogenation under reduced pressure and/or by mixing inan inert gas. Suitable inert gases are, for example, nitrogen, steam,carbon dioxide and noble gases such as He, Ne or Ar. Preference is givento diluents which are inert under the reaction conditions (i.e. diluentswhich are changed chemically to an extent of less than 5 mol %,preferably less than 3 mol % and even better less than 1 mol %). Afurther advantage of dilution with steam is generally reducedcarbonization of the dehydrogenation catalyst used according to thepresent invention and thus an increased operating life, since the steamreacts with carbon formed according to the principle of carbongasification. The ratio of steam to the hydrocarbon to be dehydrogenatedis in the range from 0 to 10 mol/mol, preferably from 0.1 to 5 mol/mol.

[0057] The process of the present invention is carried out in at leastone reaction zone with simultaneous generation of heat by exothermicreaction of hydrogen, hydrocarbon and/or carbon in the presence of anoxygen-containing gas. In general, the total amount of oxygenintroduced, based on the total amount of the hydrocarbon to bedehydrogenated, is from 0.001 to 0.5 mol/mol, preferably from 0.005 to0.2 mol/mol, particularly preferably from 0.05 to 0.2 mol/mol. Ingeneral, the amount of oxygen-containing gas added to the reaction gasmixture is chosen so that the combustion of the hydrogen or hydrocarbonpresent in the reaction gas mixture and/or the carbon present in theform of carbon deposits generates the quantity of heat required fordehydrogenation of the hydrocarbon to the alkene. In particularembodiments, the quantity of heat generated by the combustion reactionwith oxygen can also be greater or lesser than the quantity of heatrequired for the dehydrogenation of the hydrocarbon. Oxygen can be usedeither as pure oxygen or in admixture with inert gases such as CO₂, N₂or noble gases. Air is particularly preferred as oxygen-containing gas.As an alternative to molecular oxygen, it is also possible to usefurther oxygen-containing gaseous oxidants, for example dinitrogen oxideor ozone. The inert gases and the resulting combustion gases generallyhave an additional diluting effect and thus promote the heterogeneouslycatalyzed dehydrogenation.

[0058] The hydrogen burnt for heat generation can be the hydrogen formedin the dehydrogenation or be additional hydrogen added to the reactiongas mixture.

[0059] In one embodiment of the invention, no hydrogen is added to thereaction gas mixture and the heat required for dehydrogenation isgenerated at least partly by combustion (exothermic reaction) ofhydrocarbon and of the hydrogen formed in the dehydrogenation.

[0060] In a further embodiment, additional hydrogen is added to thereaction gas mixture.

[0061] The dehydrogenation catalyst used according to the presentinvention generally also catalyzes the combustion of hydrocarbons and ofhydrogen with oxygen, so that no specific oxidation catalyst differentfrom this is necessary in principle. In one embodiment, a specific,different oxidation catalyst which selectively carbonizes the oxidationof hydrogen for the generation of heat is used in addition to thedehydrogenation catalyst, particularly when additional hydrogen isadded.

[0062] If, as in one embodiment of the invention, no additional hydrogenis added to the reaction gas mixture, the heat of dehydrogenation canreadily be generated by catalytic combustion of the hydrocarbons and ofhydrogen formed in the dehydrogenation over the dehydrogenationcatalyst. Suitable, oxygen-insensitive dehydrogenation catalysts whichcatalyze the combustion of hydrocarbons are the above-describedLewis-acid catalysts. Preference is given to dehydrogenation catalystsof the type described above which comprise at least one element oftransition group VIII, at least one element of main groups I and/or II,at least one element of main groups III and/or IV and at least oneelement of transition group III including the lanthanides and actinideson zirconium oxide and/or silicon dioxide as support.

[0063] In a preferred embodiment, hydrogen is added to the reaction gasmixture for direct heat generation by combustion. In general, the amountof hydrogen added to the reaction gas mixture is such that the molarratio of H₂/O₂ in the reaction gas mixture immediately after theaddition is from 0.1 to 200 mol/mol, preferably from 1 to 20 mol/mol,particularly preferably from 2 to 10 mol/mol. In the case of multistagereactors, this applies to each intermediate introduction of hydrogen andoxygen.

[0064] The combustion of hydrogen occurs catalytically. In oneembodiment of the invention, no specific oxidation catalyst differentfrom the dehydrogenation catalyst is used. In a particularly preferredembodiment, the reaction is carried out in the presence of one or moreoxidation catalysts which selectively catalyze the combustion reactionof hydrogen and oxygen in the presence of hydrocarbons. As a result, thecombustion reaction of hydrocarbons with oxygen to give CO and CO₂proceeds only to a subordinate degree, which has a significant positiveeffect on the achieved selectivities for the formation of alkenes. Thedehydrogenation catalyst and the oxidation catalyst are preferablypresent in different reaction zones.

[0065] In a multistage reaction, the oxidation catalyst can be presentin only one reaction zone, in a plurality of reaction zones or in allreaction zones.

[0066] The catalyst which selectively catalyzes the oxidation ofhydrogen in the presence of hydrocarbons is preferably located at pointsat which higher oxygen partial pressures prevail than at other points ofthe reactor, in particular in the vicinity of the feed point for theoxygen-containing gas. Oxygen-containing gas and/or hydrogen can be fedin at one or more points on the reactor.

[0067] A preferred catalyst for the selective combustion of hydrogencomprises oxides or phosphates selected from the group consisting of theoxides and phosphates of germanium, tin, lead, arsenic, antimony andbismuth.

[0068] A further, preferred catalyst for the catalytic combustion ofhydrogen comprises a noble metal of transition group VIII or I.

[0069] In heterogeneously catalyzed dehydrogenations of hydrocarbons,small amounts of high-boiling, high molecular weight organic compoundsor carbon are generally formed over time, and these deposit on thecatalyst surface and deactivate the catalyst as time goes on. Thedehydrogenation catalysts used according to the present invention have alow tendency to suffer from carbonization and have a low deactivationrate.

[0070] The dehydrogenation catalysts used according to the presentinvention make it possible to achieve high space-time yields which forthe dehydrogenation of propane are above 2 kg of propene/kg ofcatalyst*h and are thus significantly above the space-time yields of theprocesses of the prior art. Diluting the reaction gas mixture with inertgas, increasing the reaction temperature and/or lowering the reactionpressure enable the thermodynamically possible limiting conversions tobe increased so far that they are significantly above the reactionconversions sought. In this way, space-time yields of over 6 kg ofpropene/kg of catalyst*h can be achieved in the presence of thecatalysts used according to the present invention.

[0071] The space velocity (GHSV) over the catalyst in this operatingmode also referred to as high-load operation can be >8000 h⁻¹.

[0072] Regeneration of the dehydrogenation catalyst can be carried outusing methods known per se. Thus, as described above, steam can be addedto the reaction gas mixture. The deposited carbon is partly orcompletely removed under these reaction conditions according to theprinciple of carbon gasification.

[0073] As an alternative, an oxygen-containing gas can be passed at hightemperature over the catalyst bed from time to time so as to burn offthe deposited carbon.

[0074] After a prolonged period of operation, the dehydrogenationcatalyst used according to the present invention is preferablyregenerated by, at a temperature of from 300 to 600° C., frequently atfrom 350 to 500° C., firstly carrying out a flushing operation withinert gas and subsequently, in a first regeneration step, passing airdiluted with nitrogen over the catalyst bed. The space velocity over thecatalyst is preferably from 50 to 10 000 h⁻¹ and the oxygen content isfrom about 0.5 to 2% by volume. In subsequent regeneration steps, theoxygen content is gradually increased to about 20% by volume (pure air).Preference is given to carrying out from 2 to 10, particularlypreferably from 2 to 5, regeneration steps. In general, the catalyst issubsequently regenerated further using pure hydrogen or hydrogen dilutedwith an inert gas (hydrogen content >1% by volume) under otherwiseidentical conditions. All regeneration steps are preferably carried outin the presence of water vapor.

[0075] The process of the present invention can in principle be carriedout in all reactor types known from the prior art and by all operatingprocedures known from the prior art. The additional introduction ofoxygen leads to at least part of the heat of reaction or of the energyrequired for heating the reaction gas mixture being supplied by directcombustion and not having to be introduced indirectly via heatexchangers.

[0076] A comprehensive description of suitable reactor types andoperating procedures is also given in “Catalytica® Studies Division,Oxidative Dehydrogenation and Alternative Dehydrogenation Processes,Study Number 4192 OD, 1993, 430 Ferguson Drive, Mountain View, Calif.,94043-5272 U.S.A.”.

[0077] A suitable form of reactor is a fixed-bed tube or multitube(shell-and-tube) reactor. In this, the catalyst (dehydrogenationcatalyst and, if desired, specific oxidation catalyst) is present as afixed bed in a reaction tube or in a bundle of reaction tubes. Thereaction tubes are customarily heated indirectly by a gas, e.g. ahydrocarbon such as methane, being burnt in the space surrounding thereaction tubes. In such a reactor, it is advantageous to employ thisindirect form of heating only for the first about 20-30% of the lengthof the fixed bed and to heat the remaining length of the bed to therequired reaction temperature by means of the radiative heat liberatedby the indirect heating. According to the present invention, theindirect heating of the reaction gas can advantageously be coupled withthe direct heating by combustion in the reaction gas mixture. Couplingthe direct introduction of heat with the indirect introduction of heatmakes it possible to achieve approximately isothermal reactionconditions. Customary internal diameters of the reaction tubes are fromabout 10 to 15 cm. A typical shell-and-tube reactor used fordehydrogenation has from about 300 to 1000 reaction tubes. Thetemperature in the interior of the reaction tube is generally in therange from 300 to 700° C., preferably from 400 to 700° C. The operatingpressure is usually from 0.5 to 8 bar, frequently from 1 to 2 bar whenusing low steam dilution (corresponding to the BASF Linde process) butalso from 3 to 8 bar when using a high steam dilution (corresponding tothe steam active reforming process (STAR process) of Phillips PetroleumCo., cf. U.S. Pat. No. 4,902,849, U.S. Pat. No. 4,996,387 and U.S. Pat.No. 5,389,342). In general, the product mixture leaves the reaction tubeat a from 50 to 100° C. lower temperature. Typical space velocities overthe catalyst in the case of propane are from 500 to 2000 h⁻¹. Thecatalyst geometry can be, for example, spherical or cylindrical (hollowor solid).

[0078] The process of the present invention can also be carried out in amoving bed reactor. For example, the moving catalyst bed can beaccommodated in a radial flow reactor. In this, the catalyst slowlymoves from the top downward, while the reaction gas mixture flowsradially. This method of operation is employed, for example, in theUOP-Oleflex dehydrogenation process. Since the reactors are operatedpseudoadiabatically in this process, it is advantageous to employ aplurality of reactors connected in series (typically up to fourreactors). Before or in each reactor, the inflowing gas mixture isheated to the required reaction temperature by combustion in thepresence of the oxygen fed in. The use of a plurality of reactors makesit possible to avoid large differences in the temperatures of thereaction gas mixture between reactor inlet and reactor outlet whilenevertheless achieving high total conversions.

[0079] When the catalyst has left the moving bed reactor, it is passedto regeneration and subsequently reused. The dehydrogenation catalystused according to the present invention generally has a spherical shape.Hydrogen can also be added to the hydrocarbon to be dehydrogenated,preferably propane, to avoid rapid catalyst deactivation. The operatingpressure is typically from 2 to 5 bar. The molar ratio of hydrogen topropane is preferably from 0.1 to 10. The reaction temperatures arepreferably from 550 to 660° C.

[0080] A hydrocarbon dehydrogenation by the process of the presentinvention can also, as described in Chem. Eng. Sci. 1992 b, 47 (9-11)2313, be carried out in the presence of a heterogeneous catalyst in afluidized bed, with the hydrocarbon not being diluted. In this method,two fluidized beds are advantageously operated in parallel, with one ofthem generally being in the regeneration mode. The operating pressure istypically from 1 to 2 bar, and the dehydrogenation temperature isgenerally from 550 to 600° C. The heat necessary for the dehydrogenationis introduced into the reaction system by preheating the dehydrogenationcatalyst to the reaction temperature. The use according to the presentinvention of an oxygen-containing co-feed makes it possible to omit thepreheater and to generate the necessary heat directly in the reactorsystem by combustion in the presence of oxygen.

[0081] In a particularly preferred embodiment of the process of thepresent invention, the dehydrogenation is carried out in a tray reactor.This contains one or more successive catalyst beds. The number ofcatalyst beds can be from 1 to 20, advantageously from 2 to 8, inparticular from 4 to 6. The reaction gas preferably flows radially oraxially through the catalyst beds. In general, such a tray reactor isoperated using a fixed catalyst bed.

[0082] In the simplest case, the fixed catalyst beds are arrangedaxially or in the annular gaps of concentric, upright mesh cylinders ina shaft furnace reactor. One shaft furnace reactor corresponds to onetray. It is possible for the process of the present invention to becarried out in a single shaft furnace reactor, but this is lesspreferred.

[0083] In an operating mode without oxygen as co-feed, the reaction gasmixture is subjected to intermediate heating on its way from onecatalyst bed to the next catalyst bed in the tray reactor, e.g. bypassing it over heat exchanger ribs heated by means of hot gases or bypassing it through tubes heated by means of hot combustion gases.

[0084] In the process of the present invention, the above-describedintermediate heating is carried out at least partly by direct means. Forthis purpose, a limited amount of molecular oxygen is added to thereaction gas mixture either before it flows through the first catalystbed and/or between the subsequent catalyst beds. Thus, hydrocarbonspresent in the reaction gas mixture, carbon or carbon-like compoundswhich have deposited on the catalyst surface and also hydrogen formedduring the dehydrogenation are burnt to a limited extent over thecatalyst used according to the present invention. The heat of reactionliberated in this combustion thus makes it possible for theheterogeneously catalyzed hydrocarbon dehydrogenation to be operatedvirtually isothermally. The process can be operated with or withoutintroduction of additional hydrogen.

[0085] In one embodiment of the invention, intermediate introduction ofoxygen-containing gas and possibly hydrogen is carried out upstream ofeach tray of the tray reactor. In a further embodiment of the process ofthe present invention, the introduction of oxygen-containing gas andpossibly hydrogen is carried out upstream of each tray apart from thefirst tray. In a preferred embodiment, intermediate introduction ofhydrogen is employed; in a specific embodiment of this, a bed of aspecific oxidation catalyst is present downstream of each introductionpoint and is followed by a bed of the dehydrogenation catalyst, and in asecond specific embodiment, no specific oxidation catalyst is present.In a further preferred embodiment, no hydrogen is introduced.

[0086] The dehydrogenation temperature is generally from 400 to 800° C.and the pressure is generally from 0.2 to 5 bar, preferably from 0.5 to2 bar, particularly preferably from 1 to 1.5 bar. The space velocity(GHSV) is generally from 500 to 2000 h⁻¹ and in high-load operation upto 16000 h⁻¹, preferably from 4000 to 16000 h⁻¹.

[0087] The hydrocarbon used in the process of the present invention doesnot have to be a pure compound. Rather, the hydrocarbon used cancomprise other dehydrogenatable gases such as methane, ethane, ethylene,propane, propene, butanes, butenes, propyne, acetylene, H₂S or pentanes.In particular, the dehydrogenation of the present invention can also becarried out using alkane mixtures which are produced industrially andare available in large quantities, for example LPG (liquefied petroleumgas). It is also possible to use circulation gases originating fromother processes, for example as described in the German PatentApplication P 10028582.1.

[0088] The output from the reactor is worked up in a manner known perse, for example by separating off the molecular hydrogen present in theproduct mixture, separating off constituents other than alkanes andalkenes, preferably by selective absorption of the alkene/alkane mixturein an organic solvent, and fractionation of the alkene/alkane mixture ina C₃ splitter and recirculation of the alkane to the dehydrogenation.

[0089] The invention is illustrated by the following examples.

Examples Example 1

[0090] A solution of 11.992 g of SnCl₂.2H₂O and 7.888 g of H₂PtCl₆.6H₂Oin 5950 ml of ethanol was poured over 1000 g of a granulated ZrO₂.SiO₂mixed oxide from Norton (screen fraction: 1.6 to 2 mm).

[0091] The supernatant ethanol was taken off on a rotary evaporator. Themixed oxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 7.68 g of CsNO₃, 13.54 gof KNO₃ and 98.329 g of La(NO₃)₃.6H₂O in 23 ml of H₂O was then pouredover the catalyst obtained. The supernatant water was taken off on arotary evaporator. The material was subsequently dried at 100° C. for 15hours and calcined at 560° C. for 3 hours.

[0092] The catalyst had a BET surface area of 85 m²/g. Mercuryporosimetry measurements indicated a pore volume of 0.29 ml/g.

Example 2

[0093] A solution of 0.6 g of SnCl₂.2H₂O and 0.394 g of H₂PtCl₆.6H₂O in300 ml of ethanol was poured over 55 g of a granulated ZrO₂.SiO₂ mixedoxide from Norton (screen fraction: 1.6 to 2 mm).

[0094] The supernatant ethanol was taken off on a rotary evaporator. Themixed oxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.386 g of CsNO₃, 0.680 gof KNO₃ and 4.888 g of Ce(NO₃)₃.6H₂O in 130 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 1 00° C. for 15 hoursand calcined at 560° C. for 3 hours.

[0095] The catalyst had a BET surface area of 72.4 m²/g. Mercuryporosimetry measurements indicated a pore volume of 0.26 ml/g.

Example 3

[0096] A solution of 0.684 g of SnCl₂.2H₂O and 0.45 g of H₂PtCl₆.6H₂O in342 ml of ethanol was poured over 57 g of a granulated ZrO₂ support fromNorton (screen fraction: 1.6 to 2 mm).

[0097] The supernatant ethanol was taken off on a rotary evaporator. Themixed oxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.44 g of CsNO₃, 0.775 gof KNO₃ and 5.604 g of Ce(NO₃)₃.6H₂O in 148 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

[0098] The catalyst had a BET surface area of 40 m²/g. Mercuryporosimetry measurements indicated a pore volume of 0.25 ml/g.

Example 4

[0099] In a 5 l stirred flask, 521.3 g of Zr(OH)₄ were suspended in 2000ml of H2O. 73.53 g of SiO₂ Ludox sol (SiO₂ content: 47.6% by weight)were added to the suspension. The suspension was stirred at roomtemperature for 4 hours. The product was subsequently spray dried. Thetemperature at the top was set to 350° C., the outlet temperature wasfrom 105 to 110° C., and the spraying pressure was 2.5 bar. The atomizerdisk rotated at a speed of 28 000 rpm. The resulting white powder had aloss on ignition of 15.1%.

[0100] 471.15 g of the white powder were kneaded with 133.30 g of PuralSCF (Al₂O₃) and 30.22 g of concentrated HNO₃ for 2 hours. The paste wasshaped to form 3 mm extrudates by means of a ram extruder (pressingpressure: 75 bar). The extrudates were dried at 200° C. for 4 hours andcalcined at 600° C. for 2 hours. The extrudates were subsequentlycrushed to give particles of a screen fraction from 1.6 to 2 mm.

[0101] A solution of 0.712 g of SnCl₂.2H₂O and 0.468 g of H₂PtCl₆.6H₂Oin 368 ml of ethanol was poured over 60 g of the support produced inthis way.

[0102] The supernatant ethanol was taken off on a rotary evaporator. Themixed oxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.458 g of CsNO₃, 0.807 gof KNO₃ and 5.838 g of La(NO₃)₃.6H₂O in 157 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

[0103] The catalyst had a BET surface area of 98 m²/g. Mercuryporosimetry measurements indicated a pore volume of 0.35 ml/g.

Example 5

[0104] 23 g of a granulated Mg(Al)O support from Giulini (screenfraction: 1.6 to 2 mm) were calcined at 700° C. for 2 hours. A solutionof 0.276 g of SnCl₂.2H₂O and 0.181 g of H₂PtCl₆.6H₂O in 138 ml ofethanol was subsequently poured over the support.

[0105] The supernatant ethanol was taken off on a rotary evaporator. Themixed oxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.177 g of CsNO₃, 0.313 gof KNO₃ and 2.262 g of La(NO₃)₃.6H₂O in 60 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

[0106] The catalyst had a BET surface area of 103 m²/g. Mercuryporosimetry measurements indicated a pore volume of 0.51 ml/g.

Example 6

[0107] A solution of 0.3718 g of SnCl₂.2H₂O and 0.245 g of H₂PtCl₆.6H₂Oin 190 ml of ethanol was poured over 57 g of a granulated theta-Al₂O₃support from Condea (screen fraction: 1.6 to 2 mm).

[0108] The supernatant ethanol was taken off on a rotary evaporator. Themixed oxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.239 g of CsNO₃, 0.4214g of KNO₃ and 5.604 g of La(NO₃)₃.6H₂O in 80 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

[0109] The catalyst had a BET surface area of 119 m²/g. Mercuryporosimetry measurements indicated a pore volume of 0.66 ml/g.

Example 7

[0110] A solution of 0.2758 g of SnCl₂.2H₂O and 0.1814 g of H₂PtCl₆.6H₂Oin 138 ml of ethanol was poured over 23 g of a granulated theta-Al₂O₃support from BASF (screen fraction: 1.6 to 2 mm).

[0111] The supernatant ethanol was taken off on a rotary evaporator. Themixed oxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.1773 g of CsNO₃, 0.3127g of KNO₃ and 2.26 g of La(NO₃)₃.6H₂O in 60 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

[0112] The catalyst had a BET surface area of 34 m²/g. Mercuryporosimetry measurements indicated a pore volume of 0.23 ml/g.

[0113] Example 8

[0114] 43.25 g of (NH₄)₂CO₃ were dissolved in 1 l of water, mixed with849 ml of a 25% strength by weight ammonia solution and heated to 75° C.2333.3 g of Mg(NO₃)₂.6H₂O and 337.6 g of Al(NO₃)₃.9H₂O dissolved in 3 lof water were added quickly to the solution from a dropping funnel whilestirring. After the mixture had been stirred at 75° C. for 1 hour, theresulting precipitate was filtered off and the filter cake was washedwith water. The solid was subsequently dried at 100° C. for hours andcalcined at 900° C. for 2 hours.

[0115] The powder was mixed with 3% by weight of magnesium stearate andprecompacted to form 20×2 mm tablets on an eccentric press.

[0116] 33 g of the Mg(Al)O support produced in this way were crushed(screen fraction: 1.6 to 2 mm) and calcined at 700° C. for 2 hours. Asolution of 0.398 g of SnCl₂.2H₂O and 0.262 g of H₂PtCl₆.6H₂O in 200 mlof ethanol was subsequently poured over the support.

[0117] The supernatant ethanol was taken off on a rotary evaporator. Themixed oxide granules were subsequently dried at 100° C. for 15 hours andcalcined at 560° C. for 3 hours. A solution of 0.256 g of CsNO₃, 0.451 gof KNO₃ and 3.265 g of La(NO₃)₃.6H₂O in 87 ml of H₂O was then pouredover the catalyst. The supernatant water was taken off on a rotaryevaporator. The material was subsequently dried at 100° C. for 15 hoursand calcined at 560° C. for 3 hours.

[0118] The catalyst had a BET surface area of 85 m²/g. Mercuryporosimetry measurements indicated a pore volume of 0.28 ml/g.

Example 9

[0119] Catalyst Test:

[0120] 20 ml of the previously produced catalyst were installed in atube reactor having an internal diameter of 20 mm. The catalyst wastreated with hydrogen for 30 minutes at 500° C. The catalyst was thenexposed to a mixture of 80% by volume of nitrogen and 20% by volume ofair (lean air) at the same temperature. After a flushing phase of 15minutes using pure nitrogen, the catalyst was reduced by means ofhydrogen for 30 minutes. The catalyst was then supplied with 20 standardl/h of propane (99.5% by volume) and H₂O in a molar ratio ofpropane/water vapor of 1:1 at a reaction temperature of 610° C.. Thepressure was 1.5 bar and the space velocity (GHSV) was 2000 h⁻¹. Thereaction products were analyzed by gas chromatography. The results aresummarized in the table.

[0121] The Brönsted and Lewis acidities of the catalysts produced inExamples 1 to 8 were determined by means of the probe gas pyridine usingan HV-FTIR measurement cell.

[0122] The samples were baked in air at 390° C. for 1 hour, subsequentlyevacuated to 10⁻⁵ mbar, cooled to 80° C. and treated with gaseouspyridine at an equilibrium pressure of 3 mbar. To test the evacuationstability of the pyridine adsorbate, the sample which had been treatedwith gaseous pyridine at 3 mbar was subjected to a vacuum treatment inan oil pump vacuum (about 10⁻² mbar, 3 min) and under high vacuum (about10⁻⁵ mbar, 1 h). Physisorbate material was thus desorbed. The adsorbatespectra were recorded in a high vacuum.

[0123] The measured absorbances were expressed as a ratio to thethickness of the sample (in integrated extinction units (IEE) per μm ofthickness). The single-beam spectrum of the sample which had not beentreated with pyridine gas and had been cooled to 80° C. under highvacuum served as background for the adsorbate spectra. Matrix bands werethus completely balanced out.

[0124] The band at 1440 cm⁻¹ (corresponds to Lewis-acid centers) andadditionally the control band at 1490 cm⁻¹ (corresponds to Lewis-acidcenters if no Brönsted-acid centers are present) were evaluated.

[0125] The results are summarized in the table.

[0126] All samples examined displayed no measurable Brönsted acidity.The measured Lewis acidity correlates well with the conversion in thedehydrogenation of propane. TABLE Propane Lewis acidity conversionExample Support (AU) (%) 1 ZrO₂/SiO₂ (Norton) 8.97 46.8 2 ZrO₂/SiO₂ (Ceinstead of La) 7.82 47.2 3 ZrO₂ (Norton) 4.59 30 4 ZrO₂/SiO₂/Al₂O₃ 7.2942.1 5 Mg(Al)O (Giulini) 2.88 12.4 6 theta-Al₂O₃ (Condea) 3.51 30.4 7theta-Al₂O₃ (BASF) 1.77 11.4 8 Mg(Al)O 1.66 12.2

Example 10

[0127] High-Load Operation

[0128] 2.5 ml of the catalyst produced as described in Example 1 werediluted with 77.5 ml of steatite and installed in a tube reactor havingan internal diameter of 20 mm. The catalyst was treated in succession,for 30 minutes each at 500° C., firstly with hydrogen, then with leanair (80% by volume of nitrogen and 20% by volume of air) andsubsequently again with hydrogen. The catalyst was flushed with nitrogenfor 15 minutes between each of the treatments. The catalyst wassubsequently supplied at 600° C. with 20 standard l/h of propane (99.5%by volume) and water vapor in a molar ratio of propane/H₂O of 1:1. Thepressure was 1.5 bar and the space velocity (GHSV) was 16 000 h⁻. Thereaction products were analyzed by gas chromatography. After a reactiontime of one hour, 30% of the input propane were converted at aselectivity to propene of 95%. The space-time yield of propene, based onthe catalyst volume used, was 8 g of propene/(g of catalyst*h).

Example 11

[0129] Operation with Oxygen

[0130] 20 ml of the catalyst produced as described in Example 1 wereinstalled in a tube reactor having an internal diameter of 20 mm. Thecatalyst was treated in succession, for 30 minutes each at 500° C.,firstly with hydrogen, then with lean air (80% by volume of nitrogen and20% by volume of air) and subsequently again with hydrogen. The catalystwas flushed with nitrogen for 15 minutes between each of the treatments.The catalyst was subsequently supplied at 610° C. with 20 standard l/hof propane (99.5% by volume) and water vapor in a molar ratio ofpropane/H₂O of 1:1. In addition, oxygen was introduced in a molar ratioof propane/O₂ of 20:1. The pressure was 1.5 bar and the space velocity(GHSV) was 2100 h⁻¹. The reaction products were analyzed by gaschromatography. After a reaction time of one hour, 50% of the inputpropane were converted at a selectivity to propene of 90%. After areaction time of 16 hours, the conversion was 44% and the selectivitywas 90%.

Example 12

[0131] Operation with Oxygen at Low Conversion

[0132] 20 ml of the catalyst produced as described in Example 1 wereinstalled in a tube reactor having an internal diameter of 20 mm. Thecatalyst was treated in succession, for 30 minutes each at 500° C.,firstly with hydrogen, then with lean air (80% by volume of nitrogen and20% by volume of air) and subsequently again with hydrogen. The catalystwas flushed with nitrogen for 15 minutes between each of the treatments.The catalyst was subsequently supplied at 500° C. with 20 standard l/hof propane (99.5% by volume) and water vapor in a molar ratio ofpropane/H₂O of 1:1. In addition, oxygen was introduced in a molar ratioof propane/O₂ of 20:1. The pressure was 1.5 bar and the space velocity(GHSV) was 2100 h⁻¹. The reaction products were analyzed by gaschromatography. After a reaction time of one hour, 16% of the inputpropane were converted at a selectivity to propene of 99%. After areaction time of 100 hours, the conversion was 14% and the selectivitywas 94%. After increasing the temperature to 510° C., the propaneconversion after 300 hours was 15% and the selectivity was 94%. Afterincreasing the temperature further to 530° C., 15% of the input propanewere being converted at a selectivity to propene of 94% after 800 hours.After 1700 hours, the supply of propane and of water was stopped andlean air (80% by volume of nitrogen and 20% by volume of air) was passedover the catalyst at 400° C. Pure air was then passed over the catalystfor 30 minutes. After the reactor had been flushed with nitrogen for 15minutes, hydrogen was passed over the catalyst for 30 minutes. Afterpropane, water vapor and oxygen were again supplied as feed, a propaneconversion of 15% at a selectivity of 92% could be achieved at 505° C.After a total of 2300 hours, the propane conversion at 540° C. was 15%and the selectivity to propene was 94%.

Example 13

[0133] Operation with Oxygen at Low Conversion Using Additional N₂Dilution

[0134] After 2300 hours, the catalyst from Example 1 was again (afterstopping propane and water vapor supply) treated with lean air (80% byvolume of nitrogen and 20% by volume of air) at 400° C. Subsequently,pure air was passed over the catalyst for 30 minutes. After the reactorhad been flushed with nitrogen for 15 minutes, hydrogen was passed overthe catalyst for 30 minutes. Subsequently, propane, nitrogen, oxygen andwater vapor in a ratio of 5.8/7.8/0.4/5.8 were passed over the catalystat 505° C. The reaction pressure was 1.5 bar and the space velocity(GHSV) was 1300 h³¹ ¹. The propane conversion was 20% at a selectivityof 92%. After 500 hours, 20% of the propane were converted at aselectivity to propene of 92%.

Example 14

[0135] Operation with Oxygen at Low Conversion Using Additional N₂Dilution and Additional Introduction of H₂

[0136] Using the experimental procedure of Example 12, hydrogen wasadditionally mixed into the feed after a total running time of 2500hours. The feed then had the following composition:C3/N₂/O₂/H₂/H₂O=5.8/7.8/0.4/0.8/5.8. The reaction pressure was 1.5 barand the space velocity (GHSV) was 1300 h⁻¹. The reaction temperature wasset to 575° C. The propane conversion was 20% at a propene selectivityof 92%. The oxygen introduced was reacted completely. 60% of the oxygenintroduced reacted with propane or propene to form carbon dioxide andcarbon monoxide, 40% of the oxygen introduced reacted with hydrogenwhich had been introduced or formed in the dehydrogenation to givewater.

Example 15

[0137] Regeneration of the Catalyst

[0138] 1000 ml of the catalyst produced as described in Example 6 werediluted with 500 ml of steatite and installed in a tube reactor havingan internal diameter of 40 mm. The catalyst was treated in succession,for 30 minutes each at 500° C., firstly with hydrogen, then with leanair (80% by volume of nitrogen and 20% by volume of air) andsubsequently again with hydrogen. The catalyst was flushed with nitrogenfor 15 minutes between each of the treatments. The catalyst wassubsequently supplied at 610° C. with 250 standard l/h of propane (99.5%by volume) and water vapor in a molar ratio of propane/H₂O of 1:1. Thepressure was 1.5 bar and the space velocity (GHSV) was 500 h⁻¹. Thereaction products were analyzed by gas chromatography. After a reactiontime of one hour, 55% of the input propane were converted at aselectivity to propene of 90%. After a reaction time of 12 hours, theconversion was 53% and the selectivity was 93%. The supply of propaneand water was stopped and lean air (92% by volume of nitrogen and 8% byvolume of air) was passed over the catalyst at 400° C. The air contentwas subsequently increased twice (firstly to 83% by volume of nitrogenand 17% by volume of air, then to 64% by volume of nitrogen and 36% byvolume of air). Pure air was then passed over the catalyst until the CO₂content of the output gas was less than 0.04% by volume. After thereactor had been flushed with nitrogen for 15 minutes, hydrogen waspassed over the catalyst for 30 minutes. After propane, water vapor andoxygen were again supplied as feed, a propane conversion of 55% at aselectivity of 92% could be achieved at 610° C. After the catalyst hadbeen regenerated 10 times in the above-described manner, a conversion of54% at a propene selectivity of 93% could be achieved at 610° C. Afterthe catalyst had been regenerated 30 times, a conversion of 54% at apropene selectivity of 93% could be achieved at 610° C.

We Claim:
 1. A process for the heterogeneously catalyzed dehydrogenationin one or more reaction zones of one or more dehydrogenatableC₂-C₃₀-hydrocarbons in a reaction gas mixture comprising them, with atleast part of the heat of dehydrogenation required being generateddirectly in the reaction gas mixture in at least one reaction zone bycombustion of hydrogen, the hydrocarbon or hydrocarbons and/or carbon inthe presence of an oxygen-containing gas, wherein the reaction gasmixture comprising the dehydrogenatable hydrocarbon or hydrocarbons isbrought into contact with a Lewis-acid dehydrogenation catalyst whichhas essentially no Brönsted acidity, wherein the dehydrogenationcatalyst has a Lewis acidity of greater than 3 acidity units (AU),determinable from IR absorption spectra of pyridine adsorbed on thecatalyst.
 2. A process as claimed in claim 1, wherein thedehydrogenation catalyst comprises a metal oxide selected from the groupconsisting of zirconium dioxide, aluminum oxide, silicon dioxide,titanium dioxide, magnesium oxide, lanthanum oxide and cerium oxide. 3.A process as claimed in claim 2, wherein the dehydrogenation catalystcomprises zirconium dioxide and/or silicon dioxide.
 4. A process asclaimed in any of claims 1 to 3, wherein the dehydrogenation catalystcomprises at least one element of transition group VIII, at least oneelement of main group I or II, at least one element of main group III orIV and at least one element of transition group III including thelanthanides and actinides.
 5. A process as claimed in any of claims 1 to4, wherein the dehydrogenation catalyst comprises platinum and/orpalladium.
 6. A process as claimed in any of claims 1 to 5, wherein thedehydrogenation catalyst comprises cesium and/or potassium.
 7. A processas claimed in any of claims 1 to 6, wherein the dehydrogenation catalystcomprises lanthanum and/or cerium.
 8. A process as claimed in any ofclaims 1 to 7, wherein the dehydrogenation catalyst comprises tin.
 9. Aprocess as claimed in any of claims 1 to 8, wherein the dehydrogenationcatalyst has a bimodal pore radius distribution in which from 70% to100% of the pores have a pore diameter less than 20 nm or in the rangefrom 40 to 5000 nm.
 10. A process as claimed in any of claims 1 to 9,wherein the reaction gas mixture comprises water vapor.
 11. A process asclaimed in any of claims 1 to 10, wherein hydrogen is added to thereaction gas mixture.
 12. A process as claimed in claim 11, wherein atleast one reaction zone contains a catalyst which selectively catalyzesthe combustion reaction of hydrogen and oxygen in the presence ofhydrocarbons.
 13. A process as claimed in any of claims 1 to 12, whereinthe catalyst which catalyzes the combustion of hydrogen comprises oxidesor phosphates selected from the group consisting of the oxides andphosphates of germanium, tin, lead, arsenic, antimony and bismuth.
 14. Aprocess as claimed in any of claims 1 to 13, wherein the catalyst whichselectively catalyzes the combustion of hydrogen comprises a noble metalof transition group VIII or I.
 15. A process as claimed in any of claims1 to 14, wherein the dehydrogenation is carried out in a tray reactor.